Isomerization, catalytic activation and alkylation of pentane-enriched hydrocarbon mixtures

ABSTRACT

Processes for producing liquid transportation fuels by converting a hydrocarbon feed stream comprising both isopentane and n-pentane. The hydrocarbon feed stream is separated into a first fraction that predominantly comprises isopentane and a second fraction that predominantly comprises n-pentane and some C6 paraffins. The first fraction is catalytically activated to an activation effluent comprising olefins and aromatics, while the second fraction is isomerized to convert at least a portion of the n-pentane to isopentane, then combined with the hydrocarbon feed stream to allow the newly-produced isopentane to be separated into the first fraction. At least a portion of the activation effluent is alkylated to enhanced yields of products that are suitable for use as a blend component of liquid transportation fuels.

CROSS-REFERENCE TO RELATED APPLICATIONS

This application is a non-provisional application which claims thebenefit of and priority to U.S. Provisional Application Ser. No.62/838,081 filed Apr. 24, 2019, titled “Isomerization, CatalyticActivation and Alkylation of Pentane-Enriched Hydrocarbon Mixtures,”which is hereby incorporated by reference in its entirety.

STATEMENT REGARDING FEDERALLY SPONSORED RESEARCH OR DEVELOPMENT

None.

FIELD OF THE INVENTION

The present disclosure generally relates to processes and systems thatconverts at least a portion of the n-pentane in a light hydrocarbon feedstream to isopentane, followed by an activation step and subsequentupgrading to larger hydrocarbons in either an alkylation reactor oroligomerization reactor. The processes and systems produce hydrocarbonssuitable for use as a blend component of a liquid transportation fuel.

BACKGROUND

A large surplus of pentanes are available in the petroleum refiningindustry, arising predominantly from the increased production of lighthydrocarbons from U.S. shale formations, and also from limits on thequantity of volatile components that can be blended into finishedtransportation fuels, which must adhere to regulations on maximum vaporpressure. Unfortunately, conventional processes for upgrading lightalkanes to value-added products are not well-suited for hydrocarbon feedstreams that primarily comprise pentanes (i.e., isopentane andn-pentane). Therefore, it would be beneficial to find improved processesand systems for efficiently converting pentanes to more valuableproducts, including transportation fuels and chemicals, while minimizingthe production of C1-C4 light paraffins.

The inventive processes disclosed herein provide an improved upgradingroute for pentane-rich fuel blend-stocks and other pentane-rich streamsthat do not meet government specifications for a transportation fuel.The inventive processes and systems provide enhanced yields of upgradedproducts that may be suitable for use as transportation fuels or othervalue-added chemical products.

BRIEF SUMMARY OF THE DISCLOSURE

Certain embodiments comprise a method for converting a feed streamcomprising pentanes to produce a liquid transportation fuel, the methodcomprising: a) providing a hydrocarbon feed stream comprising at least50 wt. % pentanes, including both n-pentane and isopentane; b) at leastpartially separating the hydrocarbon feed stream in a separator toproduce: a first fraction that comprises at least 80% of the isopentanepresent in the hydrocarbon feed stream (on a molar basis), and thatfurther comprises at least 90% of hydrocarbons present in thehydrocarbon feed stream (on a molar basis) that are characterized by avapor pressure equal to or greater than the vapor pressure ofisopentane, and a second fraction that comprises at least 80% of then-pentane content of the hydrocarbon feed stream (on a molar basis) andat least 90% of hydrocarbons content from the hydrocarbon feed streamthat contains six or more carbons (on a molar basis); c) contacting thefirst fraction with an activation catalyst at a temperature and apressure that facilitates catalytic activation of at least a portion ofthe first fraction by the activation catalyst to produce an activationeffluent comprising olefins containing from two to five carbon atoms,monocyclic aromatics and unconverted alkanes containing from two to fivecarbon atoms; d) removing a first portion of the second fraction for useas blend component of a liquid transportation fuel; e) contacting asecond portion the second fraction with at least one isomerizationcatalyst in an isomerization reactor that is maintained at a temperatureand a pressure that facilitates the isomerization of at least a portionof n-pentane in the second fraction to produce isopentane, therebyproducing an isomerization effluent characterized by an increased ratioof isopentane to n-pentane (on a wt. % basis) relative to the isopentaneto n-pentane ratio of the hydrocarbon feed stream; f) combining theisomerization effluent with the hydrocarbon feed stream; g) alkylatingat least a portion of the activation effluent of c) by contacting itwith an alkylation catalyst at conditions of temperature and pressurethat facilitates the alkylation of monocyclic aromatics to producealkylated monocyclic aromatics, thereby producing an alkylation effluentthat comprises an increased wt. % of mono-alkylated aromatichydrocarbons containing from seven to nine carbon atoms relative to themono-alkylated aromatic hydrocarbon content (in wt. %) of the activationeffluent.

Some embodiments further comprise separating the liquid hydrocarbonsfraction into an aromatics fraction and an unreacted C5/C6 hydrocarbonsfraction, where the aromatics fraction comprises monocyclic aromaticssuitable for use as a blend component of gasoline and the unreactedC5/C6 hydrocarbons fraction comprises alkanes and olefins containingfrom five to six carbons. In certain embodiments, the unreacted C5/C6hydrocarbons fraction is mixed with the hydrocarbon feed stream of a).

Some embodiments additionally comprise compressing and partiallycondensing the activation effluent to produce 1) a liquid hydrocarbonsfraction comprising a portion of the monocyclic aromatics present in theactivation effluent and unreacted alkanes containing at least fivecarbon atoms, and 2) a compressed alkylation feed stream comprisinguncondensed monocyclic aromatics, hydrogen and C1-C5 alkanes andolefins, wherein the compressed alkylation feed stream is subjected tothe alkylating of g).

Some embodiments additionally comprise splitting the alkylation effluentto produce a heavy hydrocarbons fraction comprising at least 80 wt. %hydrocarbons that contain at least five carbon atoms, and a lighthydrocarbons fraction comprising hydrogen and further comprising atleast 80 wt. % hydrocarbons that contain four or fewer carbon atoms.

In some embodiments, the light hydrocarbons fraction is separated toproduce a hydrogen stream and a light paraffins stream comprisingparaffins containing four or less carbon atoms. In some embodiments, theunreacted C5/C6 hydrocarbons fraction is mixed with the hydrocarbon feedstream of a). In some embodiments, the hydrocarbon feed stream comprisesat least 70 wt. % pentanes.

In some embodiments, the activation catalyst comprises one or morezeolites characterized by a Si/Al ratio ranging from 12 to 80. In someembodiments, the activation catalyst comprises ZSM-5 zeolite. In someembodiments, the activation catalyst facilitates carbon-hydrogen bondactivation that leads to at least one of olefination, dimerization,oligomerization and aromatization of the first fraction. In someembodiments, the temperature in the activation reactor is maintained ata temperature in the range from 500° C. to 625° C. and a pressure in therange from 15 psig to 100 psig. In some embodiments, temperature in therange from 550° C. to 600° C. and a pressure in the range from 20 psigto 60 psig.

Some embodiments further comprise adding a diluent to the first fractionprior to the contacting of the first fraction with the activationcatalyst, wherein the diluent inhibits catalytic conversion of the firstfraction by the activation catalyst at the conditions of temperature andpressure that are maintained in the activation reactor. Some embodimentsfurther comprise adding a diluent to the first fraction prior to thecontacting of c), wherein the diluent is characterized as chemicallynon-reactive with the activation catalyst at the conditions oftemperature and pressure that are maintained in the activation reactor.

In some embodiments, the diluent is added in an amount that alters thespecificity of the activation catalyst to increase the production ofolefins, decrease the production of aromatics, or combinations thereof,thereby increasing the ratio of olefins to aromatics in the activationeffluent. In some embodiments, the diluent is added in an amount that iseffective to produce an activation effluent that is characterized by anolefins to aromatics ratio in the range of 0.5 to 2.0.

In some embodiments, the diluent is selected from at least one ofmethane, ethane, propane and butane isomers. In some embodiments, atleast a portion of the light paraffins stream is mixed with the firstfraction and serves as the diluent.

In some embodiments, at least a portion of the second fraction is passedthrough multiple isomerization reactors that are arranged in seriesconfiguration, wherein each of the multiple isomerization reactorscontains at least one isomerization catalyst and is maintained at atemperature and a pressure that facilitates the isomerization ofn-pentane to isopentane by the at least one isomerization catalyst.

BRIEF DESCRIPTION OF THE DRAWINGS

A more complete understanding of the present invention and benefitsthereof may be acquired by referring to the follow description taken inconjunction with the accompanying drawings in which:

FIG. 1 is a diagram depicting a first embodiment of the inventiveprocesses and systems.

FIG. 2 is a diagram depicting a second embodiment of the inventiveprocesses and systems.

FIG. 3 is a diagram depicting a third embodiment of the inventiveprocesses and systems.

FIG. 4 is a bar graph depicting product selectivity resulting fromcatalytic activation of either n-pentane or iso-pentane at two differenttemperatures.

FIG. 5 is a bar graph showing the effect of isomerization of the feedstream on the total conversion and product yield for a first feed streamcomprising a 1:1 ratio of n-C5 to i-C5, and a second feed streamcomprising a 7:3 ratio of n-C5 to i-C5.

FIG. 6 is a bar graph showing the effect of isomerization of the feedstream on the total conversion and product selectivity for a first feedstream comprising a 1:1 ratio of n-C5 to i-C5, and a second feed streamcomprising a 7:3 ratio of n-C5 to i-C5.

The invention is susceptible to various modifications and alternativeforms, specific embodiments thereof are shown by way of example in thedrawings. The drawings may not be to scale. It should be understood thatthe drawings are not intended to limit the scope of the invention to theparticular embodiment illustrated.

DETAILED DESCRIPTION

The present disclosure provides processes and systems for converting amixture of light hydrocarbons to liquid transportation fuels. Morespecifically, it pertains to the conversion of any hydrocarbon mixturethat predominantly comprises pentanes to generate upgraded products thatmay be sold as a value-added chemical or utilized as a blend componentof a liquid transportation fuel.

Generally speaking, the inventive processes and systems described hereinutilize a hydrocarbon feed stream that comprises both isopentane (i-C5)and n-pentane (n-C5). The process performs an initial separation of thehydrocarbon feed stream to produce a first fraction that predominantlycomprises i-C5 and a second fraction that predominantly comprises n-C5and C6+ hydrocarbons. The first fraction is then catalytically-activatedin an activation reactor to convert at least a portion of the i-C5 toolefins and aromatics. A portion of the second fraction is eitherdiverted to blended into liquid transportation fuels or isomerized in anisomerization reactor to convert at least a portion of the n-C5 in thesecond fraction to i-C5, and a portion of any n-hexane (n-C6) in thesecond fraction to iso-hexane isomers. The resulting isomerizationeffluent is then recycled to the initial separation step to allowseparation of the additional i-C5 produced in the isomerization reactor,which is then directed to the activation reactor along with the firstfraction.

The inventive processes and systems allow an increased percentage of thehydrocarbon feed stock to be catalytically activated to producepredominantly olefins and aromatics, while minimizing the percentage ofthe feed stock that is converted to less desirable C1-C4 paraffins. Theresulting isomerization effluent is then catalytically activated underconditions of temperature and pressure (typically measured at the inletof the activation reactor) that maximize the catalytic conversion of theisomerization effluent to olefins and aromatics, while minimizing theundesirable production of C1-C4 light hydrocarbons, often referred to asfuel gas.

Certain embodiments of the inventive processes and systems furtherupgrade the resulting activation effluent by contacting the activationeffluent with an oligomerization and/or alkylation catalyst at atemperature and pressure that facilitates conversion of the activationeffluent to value-added chemicals and/or products suitable for use as aliquid transportation fuel blend component. This additional upgradingmay be preceded by compressing (and optionally, partially condensing)the activation effluent in order to facilitate the subsequentoligomerization and/or alkylation step.

As mentioned, the hydrocarbon feed stream predominantly comprises amixture of pentane isomers (C5). The hydrocarbon feed stream mayadditionally comprise C2-C4 hydrocarbons, C6-C7 hydrocarbons, or both.The hydrocarbon feed stream generally comprises both i-C5 and n-C5pentane isomers. In certain embodiments, the hydrocarbon feed streamcomprises at least 10 wt. % of a mixture of pentane isomers; optionally,at least 20 wt. %, optionally, at least 30 wt. %, optionally, at least40 wt. %, optionally at least 50 wt. %, optionally, at least 60 wt. %,or optionally, at least 70 wt. %. of a mixture of pentane isomers. Incertain embodiments, the hydrocarbon feed stream may be obtained byprocessing a stream of natural gas liquids to remove lighter components(i.e., C1-C4) by way of conventional natural gas processing technologiesthat are well-characterized, such as de-methanizer, de-ethanizer,de-propanizer and de-butanizer fractionation columns, although suchseparation is not essential. A typical result of such processes iscommonly characterized as natural gasoline, comprising about 72 wt. %pentanes, with the remainder predominantly comprises hexane. In certainembodiments, the hydrocarbon feed stream comprises less than 30 wt. %,optionally, less than 20 wt. %, optionally, less than 10 wt. % ofhydrocarbons containing four or fewer carbon atoms.

A first embodiment of the inventive processes and systems is illustratedby the process flow-diagram of FIG. 1. A hydrocarbon feed stream 101comprising both n-pentane (nC5) and isopentane (iC5) is converted in asystem 10. The hydrocarbon feed stream 101 is received by a firstseparator 110 that operates to separate the feed stream 101 into a firstfraction 119 that comprises an increased wt. % of isopentane (i-C5)relative to the feed stream 101, and a second fraction 117 thatcomprises an increased wt. % of n-pentane relative to the feed stream101. The second fraction further comprises a large majority of any C6and larger (C6+) hydrocarbons originally present in the hydrocarbon feedstream 101. The first separator 110 may operate using conventionalseparation technology 115 to separate the i-C5 isomer from the remainingcompounds present within the C5-rich hydrocarbon feed stream.Alternatively, any other separation technology may be used to assist inseparating i-C5 from n-C5 to produce the first fraction and the secondfraction.

Following separation, the first fraction 119 is conveyed to anactivation reactor 133 containing an activation catalyst 135. Theactivation reactor 133 is operable to maintain a temperature andpressure that is suitable to facilitate conversion of the first fraction119 to a first effluent 136.

Speaking generally, the temperature within the activation reactor(typically measured at, or proximal to, the inlet of the activationreactor) is maintained in the range from 500° C. to 650° C.; optionally,within the range from 525° C. to 625° C.; optionally, within the rangefrom 525° C. to 600° C.; optionally, within the range from 550° C. to600° C.; optionally, within the range from 550° C. to 575° C.;optionally, within the range from 575° C. to 600° C.

The pressure maintained within the activation reactor generally rangesfrom 15 psia (103 kpa) to 165 psia (1138 kpa); Optionally, in the rangefrom 15 psig to 100 psig.; Optionally, in the range from 15 psig to 75psig.; Optionally, in the range from 20 psig to 60 psig.; Optionally, inthe range from 20 psig to 50 psig.

Speaking generally, the activation catalyst may comprise a singlecatalyst, or a mixture of different catalysts that contact the alkanespresent in a given feed stream and facilitates at least one ofdehydrogenation, cracking, and aromatization of the alkanes, therebyproducing upgraded products including olefins and aromatics. Favoredcatalysts include supported or unsupported solid acids, metals, metalchalcogenides, or metal pnictogenides, including (but not limited to)structured and amorphous silica-aluminas, structured and amorphous solidphosphoric acids, clays, other metal oxides, metal sulfates, or metalphosphates, and graphite-supported materials. In certain embodiments,ZSM-5 zeolite catalysts are utilized that are characterized by Si/Alratios ranging from 12-80, optionally ranging from 35 to 50. Optionally,one or more elements may optionally be impregnated on the zeolitecatalyst, including one or more of Ga, Pt, Ni, Mn, Mg, Fe, Cr, P, Cu,La, Sr and F, but are not required for operability of the invention.

Generally speaking, dehydrogenation catalysts are not utilized forparaffin activation in the instant inventive process. A sufficientconcentration of intermediate olefins can be generated through acombination of thermal dehydrogenation and catalytic cracking such thattypical dehydrogenation metals can be avoided without significantlydecreasing product yield. Known dehydrogenation catalysts are prone tofouling by sulfur and nitrogen contaminants that are often present inhydrocarbon feed streams derived from petroleum, so the ability tooperate in the absence of sensitive catalytic materials is highlyadvantageous.

The inventive process generally takes advantage of the large reactivitydifference to catalytic activation of n-C5 relative to i-C5. Forexample, when contacted with a solid acid catalyst at a temperature inexcess of 550 C, the measured activation rate of i-C5 is up to four-foldthat of n-C5, even when both isomers are simultaneously present in thesame reactor. Thus, an at least partial separation of the mixed pentanesfeed stream to form a first fraction that largely comprises i-C5 and asecond fraction that retains a large majority of the n-C5 in the feedstream, which allows the iC5 to be selectively activated at conditionscomprising a lower temperature (e.g., 550° C. versus 600° C.) and apressure that decreases selectivity to C1-C4 light paraffins duringcatalytic activation, while maximizing the selectivity of the activationtoward the production of value-added products (such as olefins andaromatics).

Table 1 (below) illustrates the difference in the activation reactivityof i-C5 versus n-C5 over a microporous silica-alumina catalyst. Feedstreams comprising either 100 wt. % i-C5 or 100 wt. % n-C5 were eachcatalytically activated in separate experiments utilizing temperaturesof either 600° C. or 550° C. The conversion and product distribution fori-C5 are shown in Table 1, columns 2 and 3, while similar results forthe activation of n-C5 are shown in Table 1, columns 4 and 5. The datashows that when comparing the activation of pentane isomers, conversionof i-C5 to olefins and aromatics is possible at a temperature about 50°C. less than is required for equivalent conversion of n-C5. To be clear,we observed that activation of the i-C5 feed stream at 550° C. convertedabout the same weight percentage of the feed stream as did activation ofn-C5 at 600° C. using the same WHSV. Further, utilizing a decreasedtemperature of 550° C. for activation of the i-C5 feed streamadvantageously decreased the production of C1-C4 light paraffins from21.0% to 19.4% by increasing the product distribution toward olefinsrather than aromatic products. Thus, the ability to separate the i-C5isomer from n-C5 isomer (and any C6+ hydrocarbons), and separatelyactivate the i-C5 enriched mixture at relatively reduced temperatureresults in approximately equivalent total conversion of the hydrocarbonfeed stream, while decreasing the formation of undesired C1-C4 lightparaffins.

Referring again to the embodiment depicted in FIG. 1, the secondfraction 117 is directed to a storage vessel 120 that is operable toboth receive the second fraction, and to optionally divert at least aportion of the second fraction to produce a blending fraction 125 thatmay be blended directly into a liquid hydrocarbon transportation fuel(e.g., gasoline). The remainder of the second fraction is conveyed fromthe storage vessel 120 to a isomerization reactor 127 that contains anisomerization catalyst 130 and is operable to receive a flow of hydrogengas 131, and further operable to maintain a temperature and pressuresuitable to facilitate the conversion of the second fraction 117 to aisomerization effluent 132 comprising an increased wt. % of i-C5 and adecreased wt. % of n-C5 relative to the second fraction. In certainembodiments, the second vessel is not required, and a portion of thesecond fraction is simply diverted via a valve (not depicted) to producethe blending fraction 125. In such embodiments, the quantity of thesecond fraction that is diverted (if any) is determined by the molecularcomposition of the second fraction 117 and the quantity of the secondfraction that is to be fed to isomerization reactor 127.

Speaking generally, the isomerization reactor contains an isomerizationcatalyst and is operable to maintain a temperature and a pressure thatfacilitates the isomerization of at least a portion of the nC5 in thesecond fraction to iC5 by the isomerization catalyst. The isomerizationreaction occurring in the isomerization reactor produces anisomerization effluent that is characterized by an increased ratio ofiC5 to nC5 (relative to the corresponding ratio of the hydrocarbon feedstream).

Speaking generally, the isomerization process is designed primarily forcontinuous catalytic isomerization of the n-C5 present in the secondfraction. The process is conducted in an isomerization reactor in thepresence of an isomerization catalyst. In certain embodiments, theisomerization catalyst may comprise more than one isomerizationcatalyst. Optionally, isomerization may be conducted in more than oneisomerization reactor arranged in series configuration, with eachreactor containing a different isomerization catalyst (or mixture ofisomerization catalysts). Each isomerization reactor maintains a partialpressure of hydrogen and operating conditions of temperature andpressure in the isomerization reaction zone that facilitateisomerization by the isomerization catalyst while minimizinghydrocracking.

Ideally, the isomerization catalyst (or catalysts) facilitates theconversion of n-pentane to the higher octane-number isopentane, whileany C6 hydrocarbons present may be converted to higher octane 2-3dimethyl butane (and similar molecules). The isomerization reaction isequilibrium-limited. For this reason, any n-pentane that is notconverted on its first pass through the isomerization reactor mayoptionally be recycled to the isomerization reactor, or converted inmultiple isomerization reactors, arranged in series configuration,thereby further increasing the ratio of i-C5 to n-C5 in the product. Therelative efficiency of separation of pentane isomers by distillation ispoor. Thus, recycling may be more effectively accomplished by amolecular sieve, which selectively adsorbs n-pentane due to its smallerpore diameter relative to isopentane.

In certain embodiments, the activity of the isomerization catalyst maybe decreased in the presence of sulfur, thereby decreasing theisomerization rate and, consequently, the octane number of the finalproduct. In such embodiments, the hydrocarbon feed stream ishydrotreated to remove sulfur prior to being conveyed to theisomerization reactor.

Generally speaking, the isomerization catalyst may comprise any knownisomerization catalyst. Currently, three basic families of light naphthaisomerization catalysts are known. The first are termed super-acidiccatalysts such as, for example, chlorinated alumina catalysts withplatinum. Super acidic isomerization catalysts are highly active andhave significant activity at temperatures as low as 265° F. (130° C.)using a lower H2/HC ratio (less than 0.1 at the outlet of the reactor).However, maintaining the high acidity of these catalysts requires theaddition of a few ppm of chloriding agent to the feedstock. At the inletof the isomerization reactor, this chloriding agent reacts with hydrogento form HCl, which inhibits the loss of chloride from the catalyst.Unlike a zeolitic catalyst, the acidic sites on a super-acidic catalystare irreversibly deactivated by water. These catalysts are alsosensitive to sulfur and oxygenate contaminants, so the feed stream isgenerally hydrotreated and dried to remove residual water contamination.Commercially-available examples of chlorided-alumina catalysts include,but are not limited to, IS614A, AT-2, AT-2G, AT-10 and AT-20 (by AkzoNobel) and ATIS-2L (by Axens). Due to their chlorinated nature, theseare very sensitive to feed impurities, particularly water, elementaloxygen, sulfur, and nitrogen. When using such super-acidic catalysts,the reactor operating temperature generally ranges from 14° C. to 175°C., while the operating pressure is generally in the range from 200 psigto 600 psig, preferably in the range from 425 psig to 475 psig.

Zeolitic isomerization catalysts require a higher operating temperatureand are effective at isomerization at temperatures ranging from 220° C.to about 315° C., preferably at a temperature ranging from 230° C. to275° C. Pressures utilized for isomerization with zeolitic isomerizationcatalysts typically range from 300 psig to 550 psig with a LHSV from 0.5to 3.0 hr⁻¹. These catalysts react as bifunctional catalysts and requirehydrogen at a H2/HC ratio ranging from about 1.5 to about 3. Zeoliticcatalysts have advantages over chlorided-alumina catalysts due tozeolitic catalyst tolerance for typical catalyst poisons sulfur,oxygenates and water. Zeolitic catalysts also do not require theinjection of a chloriding agent in order to maintain catalyst activity.

A third type of conventional isomerization catalyst that may be usefulin certain embodiments comprises sulfated zirconia/metal oxidecatalysts. These catalysts are active at relatively low temperatures(e.g., 100° C.) with the advantage of providing enhanced isoparaffinyield. Their biggest drawback is their relative sensitivity to catalystpoisons, especially water. Certainly, other examples of isomerizationcatalysts that are suitable for use with the present processes andsystems described herein are known by those having experience in thefield, and thus, require no further disclosure here.

Again, referring to the embodiment disclosed in FIG. 1, theisomerization effluent 132 is conveyed via conduit to be combined withthe hydrocarbon feed stream 101 at a location that is upstream of thefirst separator 110. This allows the additional i-C5 that is produced inthe isomerization reactor 127 to be redirected to the activation reactor133 where it is catalytically activated by the activation catalyst 135.

Generally speaking, in embodiments where the hydrocarbon feed streamincludes C6 hydrocarbons, at least a portion of any n-C6 present in thesecond fraction is isomerized to i-C6 in the isomerization reactor andcomprises a portion of the isomerization effluent. This i-C6 is routedbetween the first separator and the isomerization reactor until iteventually leaves the system as a component of the blending fraction 125that is diverted from the second fraction upstream from theisomerization reactor. This serves as an additional advantage of theinventive processes and systems, since i-C6 has a significantly higheroctane rating than n-C6 and therefore is a value-added product that canbe directly blended into a liquid transportation fuel. A furtheradvantage is that any n-C6 present in the hydrocarbon feed stream ispredominantly upgraded by isomerization, which is a more cost-effectivemethod (than catalytic activation) for upgrading n-C6 to higher octaneproducts that are suitable as blend components of a liquidtransportation fuel.

Speaking generally, the activation catalyst may comprise a singlecatalyst, or a mixture of different catalysts that contacts the alkanespresent in the isomerization effluent and facilitates at least one ofdehydrogenation, cracking, and aromatization of the alkanes, therebyconverting at least a portion of hydrocarbons present in theisomerization effluent to produce the activation effluent. Moreover, theactivation effluent comprises products that may be utilized as acommodity chemical, an intermediate amenable to further catalyticupgrading, or a transportation fuel (or a component thereof).

Activation catalysts suitable for use with the processes and systemsdescribed herein may comprise any catalyst capable of cracking and/oraromatizing hydrocarbons. Favored catalysts include supported orunsupported solid acids, metals, metal chalcogenides, or metalpnictogenides, including (but not limited to) structured and amorphoussilica-aluminas, structured and amorphous solid phosphoric acids, clays,other metal oxides, metal sulfates, or metal phosphates, andgraphite-supported materials. In certain embodiments, ZSM-5 zeolitecatalysts are utilized that are characterized by Si/Al ratios rangingfrom 12-80, optionally ranging from 35 to 50. Optionally, one or moreelements may be impregnated on the zeolite catalyst, including one ormore of Ga, Pt, Ni, Mn, Mg, Fe, Cr, P, Cu, La, Sr and F.

Generally speaking, dehydrogenation is not a prerequisite for paraffinactivation in the present inventive process. A sufficient concentrationof intermediate olefins can be generated through a combination ofthermal dehydrogenation and catalytic cracking such that typicaldehydrogenation catalyst metals (such as platinum, zinc, molybdenum, orgallium) can be avoided without significantly decreasing product yield.Conventional dehydrogenation catalysts are prone to deactivation bysulfur and nitrogen contaminants that are often present in hydrocarbonfeed streams derived from petroleum, so the ability to operate in theabsence of these sensitive catalytic materials is highly advantageous tothe process.

The present inventive processes and systems take advantage of thediffering reactivity of pentane isomers to catalytic activation.Isopentane (i-C5) exhibits catalyst-dependent activation reactivity thatis typically greater than that of n-pentane (n-C5), and the optimalactivation reactor conditions for the two isomers are thereforedistinct. Experimentally, i-C5 can be activated at lower temperaturesthan n-C5, while maintaining high yields of desired products (such asolefins and aromatics) and decreasing the yield of C1-C4 paraffins (i.e.fuel gas).

For example, utilizing a solid acid activation catalyst at temperaturein excess of 550° C., the measured activation rates differ by up to 4fold in favor of i-C5, when each isomer is contacted with the samecatalyst under identical conditions (even in the same reactorsimultaneously). Thus, an initial isomerization of the hydrocarbon feedstream to increase i-C5 content, followed by activating the resultingeffluent in catalytic activation zone, maximize the yield ofvalue-added, upgraded products (such as olefins and/or aromatics).Increasing conversion of pentane isomers to i-C5 also was found tounexpectedly decrease selectivity of the activation reaction to C1-C4light gases, which typically have little value other than as fuel gas.This helps maximize the conversion of the feed to upgraded products,which is one of many advantages of the process and systems.

Table 1 (below) illustrates the difference in the activation reactivityof i-C5 versus n-C5 over a microporous silica-alumina activationcatalyst. Feed streams comprising either 100 wt. % i-C5 or 100 wt. %n-C5 were each catalytically activated in separate experiments utilizingtemperatures of either 600° C. or 550° C. The conversion and productdistribution for i-C5 are shown in Table 1, columns 2 and 3, whilesimilar results for the activation of n-C5 are shown in Table 1, columns4 and 5.

TABLE 1 Product distributions for i-C5 or n-C5 isomer feed streamsfollowing conversion by a ⅛″ extrudate consisting of 50 wt. % aluminabinder and 50 wt. % ZSM-5 zeolite. Activation was performed bycontacting the ZSM-5 catalyst with a feed stream comprising either 100wt. % of i-C5 or 100 wt % of n-C5. Results were time-averaged over 16hours and all reactions were performed at 1 atm with a WHSV = 4.0 hr⁻¹.Feed Isomer: i-C5 i-C5 n-C5 n-C5 Inlet Temperature: 600° C. 550° C. 600°C. 550° C. Conversion (wt. %): 94.5 82.4 78.5 48.3 Product Distribution(wt. %) Hydrogen 2.4% 1.6% 1.1% 0.4% Methane 9.8% 7.8% 5.3% 2.3% Ethane3.0% 2.6% 11.5%  6.6% Ethylene 17.4%  15.6%  14.3%  7.7% Propane 5.3%4.8% 10.6%  9.9% Propylene 21.2%  22.1%  16.4%  10.3%  Butane 2.9% 4.2%0.8% 1.2% Butene 8.5% 9.9% 5.8% 5.5% Isopentane 5.5% 17.6%  0.1%  0%n-Pentane  0%  0% 21.5%  51.7%  Pentene 1.2% 1.8% 0.8% 1.2% C6+ alkanes0.0% 0.3% 0.0% 0.0% Benzene 4.8% 3.3% 4.9% 1.0% Toluene 11.1%  6.0% 5.3%1.6% Xylene 6.4% 2.2% 1.3% 0.6% Ethylbenzene 0.3% 0.1% 0.1% 0.0% Coke0.2% 0.1% 0.2% 0.1%

The data indicates that when comparing the activation of pentaneisomers, conversion of i-C5 to olefins and aromatics is possible at atemperature about 50° C. less than is required for equivalent conversionof n-C5. To be clear, we observed that activation of the i-C5 feedstream at 550° C. converted about the same weight percentage of the feedstream as did activation of n-C5 at 600° C. using the same WHSV.Further, utilizing a decreased temperature of 550° C. for activation ofthe i-C5 feed stream advantageously decreased the production of C1-C4light paraffins from 21.0% to 19.4% by increasing the productdistribution toward olefins rather than aromatic products. Thus, theability to separate the i-C5 isomer from n-C5 isomer (and any C6+hydrocarbons) and activate the i-C5 enriched mixture at relativelyreduced temperature, results in approximately equivalent totalconversion of the overall feed stream, while decreasing the formation ofundesired C1-C4 light paraffins.

The inventive systems and processes take advantage of this reactivitydifference between C5 isomers by first isolating, thencatalytically-activating predominantly the i-C5 portion of a hydrocarbonfeed stream, while simultaneously isomerizing the n-C5/C6+ portion(second fraction) of the feed stream in a first isomerization step. Theisomerization effluent (that contains an increased wt. % of i-C5relative to the second fraction. This maximizes both the conversionyield and selectivity of the activation step to form useful products,including (but not limited to) olefins and aromatics, while minimizingthe selectivity of the process to form fuel gas. An additional advantageis that any C6+ hydrocarbons present in the feed stream arepredominantly isomerized to produce a value-added liquid transportationfuel blend stock without the need for additional upgrading (e.g.,catalytic activation, and optionally, oligomerization or alkylation).Additional advantages will become evident from the detailed disclosureprovided below.

Referring again to the embodiment depicted in FIG. 1, the activationeffluent 136 is conveyed into a second separator 150 that separates alight hydrocarbons 151 containing four or fewer carbons and hydrogenfrom a heavy hydrocarbons fraction 152 that predominantly comprises C5olefins, single-ring aromatics as well as unreacted pentanes and largerC6+ components originally present in the hydrocarbon feed stream 101. Incertain embodiments, the second separator 150 is a two-phase splitterand separation of the activation effluent 136 is achieved by partialcondensation.

Again, referring to the embodiment depicted in FIG. 1, lighthydrocarbons 151 predominantly comprises hydrogen as well as C1-C4hydrocarbons that were not converted in the activation reactor 133.Light hydrocarbons fraction 151 is conveyed to a fourth separator 153that typically utilizes a conventional separation technology (such as,but not limited to, pressure swing adsorption technology, membraneseparation technology, etc.) to separate hydrogen from lighthydrocarbons fraction 151 to produce a hydrogen stream 154 and a lightparaffins stream 155 that predominantly comprises paraffins containingfour or less carbons atoms.

The light paraffins stream 155 is optionally conveyed and mixed with thefirst fraction to serve as a diluent during catalytic activation of thefirst fraction 119. In such embodiments, the light paraffins stream 155is mixed with the first fraction 119 at a point that is upstream fromthe point where the first fraction 119 contacts the activation catalyst135 in the activation reactor 133.

In the embodiment depicted in FIG. 1, the heavy hydrocarbons fraction152 is conveyed to a third separator 160 that is optionally a naphthastabilizer. The third separator 160 is operable to remove an unreactedC5/C6 fraction 172 comprising predominantly alkanes and olefinscontaining five to six carbon atoms from the heavy hydrocarbons fraction152 in order to decrease Reid vapor pressure and increase octane ratingof the resulting liquid hydrocarbon product 163, which predominantlycomprises hydrocarbon molecules (predominantly monocyclic aromatics)that are characterized by a boiling-point in the range of a liquidtransportation fuel, such as, but not limited to, gasoline, diesel andjet fuel. The unreacted C5/C6 fraction 172 may be used directly as ablend component 168 of a liquid transportation fuel, may be mixed withhydrocarbon feed stream 101 at a point that is upstream from firstseparator 110, or may be mixed with the first fraction 119 (notdepicted).

Certain embodiments of the inventive processes and systems convey anactivation effluent to an oligomerization reactor containing at leastone oligomerization catalyst. The activation effluent contacts theoligomerization catalyst and is converted to larger hydrocarbon productsthat can be utilized as a component of a liquid transportation fuel,such as, but not limited to: gasoline, diesel and jet fuel. A secondembodiment of the inventive processes and systems that includes anoligomerization reactor and additional inventive features is illustratedby the process flow-diagram of FIG. 2. A hydrocarbon feed stream 201comprising both n-pentane (nC5) and isopentane (iC5) is converted in asystem 20. The hydrocarbon feed stream 201 is received by a firstseparator 210 that operates to separate the feed stream 201 into a firstfraction 219 that comprises an increased wt. % of isopentane (i-C5)relative to the hydrocarbon feed stream 201, and a second fraction 217that comprises an increased wt. % of n-pentane relative to thehydrocarbon feed stream 201. The second fraction 217 further comprises alarge majority of any C6 and larger (C6+) hydrocarbons that wereoriginally present in the hydrocarbon feed stream 201. The firstseparator 210 is operable to separate the i-C5 isomer from the remainingcompounds present within the C5-rich hydrocarbon feed stream usingconventional separation technology (e.g., distillation). Such methodsare conventional, well-understood in the field, and will therefore notbe discussed further here. Any separation technology may be used toassist in separating i-C5 from n-C5 to produce the first fraction andthe second fraction, and such separation is not considered critical tothe inventive process and systems.

Following separation, the first fraction 219 is conveyed to anactivation reactor 233 containing an activation catalyst 235. Theactivation reactor 233 is operable to maintain a temperature andpressure that is suitable to facilitate conversion of the first fraction219 to an activation effluent 236.

The activation effluent 236 leaves the activation reactor 233, and isconveyed to compressor/condenser 237, which may comprise one or morefunctions including a condenser, splitter, compressor and pump.Compressor/condenser 237 is operable to receive and condense at least aportion of the activation effluent 236 to produce a liquid hydrocarbonsfraction 238 comprising C6 and larger hydrocarbons including paraffins,olefins and aromatics. The liquid hydrocarbons fraction 238 is removedand may be used as blend stock for a liquid hydrocarbon transportationfuel, while the remaining fraction of the activation effluent 236 iscompressed by the compressor/condenser 237 to produce a compressed lighthydrocarbons fraction 239 predominantly comprising C1-C5 alkanes andolefins that remains in gaseous phase. In certain embodiments, thecompressed light hydrocarbons fraction comprises at least 50 wt. %, atleast 60 wt. %, at least 70 wt. %, or at least 80 wt. % of alkanes andolefins containing five or less carbon atoms. The compressed lighthydrocarbons fraction 239 is next conveyed to an oligomerization reactor240 that contains an oligomerization catalyst 245. The compressed lighthydrocarbons fraction 239 contacts and is catalytically transformed bythe oligomerization catalyst 245 to produce an oligomerization effluent247 that comprises an increased quantity of hydrocarbon molecules thatare characterized by a boiling-point in the range of a liquidtransportation fuel (e.g., gasoline and diesel).

Speaking generally, the oligomerization catalyst may comprise any solidcatalyst (or mixture of catalysts) characterized as possessing eitherBronsted or Lewis acidic properties. In certain embodiments, theoligomerization catalyst is a zeolite or mixture of zeolites, or areactive transition metal oxide. In certain embodiments, theoligomerization catalyst is ZSM-5, although many zeolites arewell-characterized as possessing oligomerization properties and may besuitable for use (either alone or in combination) with the inventiveprocesses and systems described herein. Other well-characterizedoligomerization catalysts include, but are not limited to: nickeloxides, aluminum alkyls, aluminum halides, perfluoroaryl boranes,oligomeric methyl aluminoxanes (including supported), perfluoroarylboranes, fluoroarylanes, trityl borate, ammonium borate (and aluminatesalts thereof), supported PhNMe2H+B(C6F5)4- and borate anions andsuperacidic solid Bronsted acids, among others.

Speaking generally, the oligomerization reactor is maintained at atemperature and pressure suitable to facilitate oligomerization ofolefins present in the gaseous activation effluent, thereby producinglarger hydrocarbons comprising at least six carbons that are preferablycharacterized by a boiling point that is in the boiling point range of aliquid transportation fuel (e.g., gasoline or diesel). Theoligomerization reactor is generally maintained at a total pressure in arange from 14 psia to 800 psia, optionally in the range from 50 psia to300 psia. The oligomerization reactor is typically maintained at atemperature (measured within the oligomerization reactor inlet) in therange from 200° C. to 420° C., optionally in the range from 200° C. to350° C. Typically, flow thorough the oligomerization reactor ismaintained at a weight hourly space velocity (WHSV) in the range from0.5 hr⁻¹ to 10 hr⁻¹. Optionally, the WHSV is in the range from 0.5 hr⁻¹to 2.0 hr⁻¹. While higher overall throughput is desirable, ideally thechosen WHSV allows for conversion of at least 85% of olefins present inthe gaseous activation effluent at the selected operating temperatureand pressure.

The catalytic conversion occurring in the oligomerization reactorproduces an oligomerization effluent that typically comprises anincreased quantity of hydrocarbon molecules that are characterized by aboiling-point in the range of a liquid transportation fuel (e.g.,gasoline and diesel). Preferably, the combination of isomerization,activation and oligomerization converts at least 30 wt. % of theoriginal hydrocarbon feed stream to hydrocarbon molecules that arecharacterized by a boiling point that is in the boiling point range ofgasoline.

Referring again to the embodiment depicted in FIG. 2, theoligomerization effluent 247 produced in the oligomerization reactor 240is conveyed to a second separator 250 that separates the oligomerizationeffluent 247 into two fractions: a light hydrocarbons fraction 251comprising C1-C4 hydrocarbons and hydrogen, and a heavy hydrocarbonsfraction 252 comprising hydrocarbons containing at least five carbonatoms (C5+) that may be utilized directly as a blend component of aliquid transportation fuel or an intermediate product that may beadditionally processed prior to blending into a liquid transportationfuel.

Again, referring to the embodiment depicted in FIG. 2, lighthydrocarbons fraction 251 predominantly comprises hydrogen as well asC1-C4 hydrocarbons that remained unconverted in the oligomerizationreactor 240. Light hydrocarbons fraction 251 leaves the second separator250 and is optionally conveyed to a fourth separator 253 that utilizes aconventional separation technology (such as, but not limited to,pressure swing adsorption technology, membrane separation technology,etc.) to separate hydrogen from the light hydrocarbons to produce ahydrogen stream 254 and a light paraffins stream 255 comprisingparaffins containing four or less carbon atoms.

The light paraffins stream 255 is optionally recycled to serve as adiluent during catalytic activation of the first fraction 219. In suchembodiments, the light paraffins stream 255 is mixed with the firstfraction 219 at a point that is upstream from the point where the firstfraction 219 contacts the activation catalyst 235 in the activationreactor 233.

The heavy hydrocarbons fraction 252 is conveyed to a third separator 260that is optionally a naphtha stabilizer. The third separator 260 isoperable to remove an unreacted C5/C6 fraction 272 comprisingpredominantly alkanes and olefins containing five to six carbon atomsfrom the heavy hydrocarbons fraction 252 in order to decrease Reid vaporpressure and increase octane rating of the resulting liquid hydrocarbonproduct 263, which predominantly comprises hydrocarbon molecules(predominantly, monocyclic aromatics) that are characterized by aboiling-point in the range of a liquid transportation fuel, such as, butnot limited to, gasoline, diesel and jet fuel. The unreacted C5/C6fraction 272 may be used directly as a blend component 268 of a liquidtransportation fuel, or alternatively, is mixed with hydrocarbon feedstream 201 at a point that is upstream from separator 210. Optionally,at least a portion of the blending fraction 225 as well as the liquidhydrocarbons fraction 238 derived from the condenser/compressor 237 maybe combined with the liquid hydrocarbon product 263.

Speaking generally, in certain embodiments, the liquid hydrocarbonproduct of the process may be hydrotreated in a hydrotreating reactorcontaining a hydrotreating catalyst in order to reduce olefin andaromatic content in the liquid hydrocarbon product, as well as to removeany nitrogen-containing and sulfur-containing compounds that may bepresent. The hydrotreating reactor contains at least one hydrotreatingcatalyst (such as, for example, NiMo, CoMo, etc.) or a precious metalcatalyst (such as Pt/Al₂O₃, Pd/Al₂O₃, or Pd/C, etc) and is maintained ata pressure and temperature suitable for facilitating hydrotreatingcatalytic reactions. Such processes are conventional in nature,well-understood in the field, and therefore will not be described ingreater detail here.

Continuing to refer to the embodiment depicted in FIG. 2, the secondfraction 217 is directed to a storage vessel 220 that is operable toboth receive the second fraction 217, and to optionally divert at leasta portion of the second fraction 217 to produce a blending fraction 225that may be blended directly into a liquid hydrocarbon transportationfuel (e.g., gasoline). Diverting may be accomplished in a number ofconventional ways (e.g., an outlet comprising an adjustable valve) thatare well-understood in the art, and thus, will not be discussed furtherhere.

The remainder of the second fraction 217 passes through the storagevessel 220 and is next conveyed to a isomerization reactor 227 thatcontains an isomerization catalyst 230 and is operable to receive a flowof hydrogen gas 231, and further operable to maintain a temperature andpressure that are suitable to facilitate the conversion of the secondfraction 217 by the isomerization catalyst 230 to produce anisomerization effluent 232 that comprises an increased wt. % of i-C5 anda decreased wt. % of n-C5 relative to the second fraction 217.

Speaking generally, certain embodiments of the inventive system andprocess do not comprise a storage vessel, and a portion of the secondfraction is simply diverted via a valve (not depicted) to produce theblending fraction. In such embodiments, the quantity of the secondfraction that is diverted (if any) is determined by the molecularcomposition of the second fraction and the percentage of the secondfraction that is to be fed to the isomerization reactor.

Certain embodiments of the inventive processes and systems convey theactivation effluent to an aromatic alkylation reactor containing atleast one alkylation catalyst. This produces larger hydrocarbon productsthat can be utilized as either gasoline or diesel transportation fuel,or a component thereof. A third embodiment of the inventive processesand systems that includes an alkylation reactor and additional inventivefeatures is illustrated by the process flow-diagram of FIG. 3. Ahydrocarbon feed stream 301 comprising both n-pentane (nC5) andisopentane (iC5) is converted in a system 30. The hydrocarbon feedstream 301 is received by a first separator 310 that operates toseparate the feed stream 301 into a first fraction 319 that comprises anincreased wt. % of isopentane (i-C5) relative to the hydrocarbon feedstream 301, and a second fraction 317 that comprises an increased wt. %of n-pentane relative to the hydrocarbon feed stream 301. The secondfraction 317 further comprises a large majority of any C6 and larger(C6+) hydrocarbons that were originally present in the hydrocarbon feedstream 301. The first separator 310 is operable to separate the i-C5isomer from the remaining compounds present within the C5-richhydrocarbon feed stream using conventional separation technology (e.g.,distillation). Such methods are conventional, well-understood in thefield, and will therefore not be discussed further here. Any separationtechnology may be used to assist in separating i-C5 from n-C5 to producethe first fraction and the second fraction, and such separation is notconsidered critical to the inventive process and systems.

Following separation, the first fraction 319 is conveyed to anactivation reactor 333 containing an activation catalyst 335. Theactivation reactor 333 is operable to maintain a temperature andpressure that is suitable to facilitate conversion of the first fraction319 to an activation effluent 336. The activation effluent 336 leavesthe activation reactor 333, and is conveyed to compressor/condenser 337,which may comprise one or more functions including condensing,splitting, compressing and pumping. Optionally, compressor/condenser 337is operable to receive and condense a portion of the activation effluent336 to produce a liquid hydrocarbons fraction 338 comprising C6 andlarger hydrocarbons including paraffins, olefins and aromatics. In suchembodiments, the liquid hydrocarbons fraction 338 is removed from thesystem via an outlet in the compressor/condenser 337 and is has thephysical and chemical characteristics meeting requirements for a liquidhydrocarbon transportation fuel, or a component thereof. Optionally, thecompressor/condenser 337 does not operate to condense the activationeffluent. In such embodiments, the entire activation effluent 336 iscompressed by compressor/condenser 337.

Following the optional partial condensation of the activation effluent336, the remaining fraction of the activation effluent 336 is compressedby the compressor/condenser 337 to produce a compressed alkylation feedstream 339 comprising gas-phase aromatics (that were not condensed bythe compressor/condenser 337) as well as C1-C5 alkanes and olefins. Thecompressed alkylation feed stream 339 is next conveyed to an alkylationreactor 340 that contains an alkylation catalyst 345. Note that thecompressor/condenser 337 does not condense all aromatics present in theactivation effluent 336, but instead retains at least a portion of thearomatics to subsequently be alkylated by olefins in the alkylationreactor 340. In alkylation reactor 340 the compressed alkylation feedstream 339 contacts the alkylation catalyst, producing an alkylationeffluent that comprises an increased wt. % of alkylated aromatics,especially monocyclic alkylated aromatics.

Speaking generally, the alkylation reactor is maintained at a feed inlettemperature and a pressure that are suitable to facilitate the catalyticalkylation of aromatics that are present in the compressed alkylationfeed stream. The aromatics that are alkylated may be produced byaromatization that takes place in the activation reactor or may be aconstituent of the hydrocarbon feed stream. Olefins that are used toalkylate the aromatics are mostly produced by the activation of alkanesin the activation reactor. Alkylation of aromatics with light olefins inthe alkylation reactor produces an alkylation effluent comprising largerhydrocarbons comprising at least seven carbons that are preferablycharacterized by a boiling point that is in the boiling point range of aliquid transportation fuel (e.g., gasoline or diesel). Typically, thealkylation effluent comprises an increased percentage of alkylatedaromatic compounds comprising from seven to nine carbon atoms.Optionally, the larger hydrocarbons also are characterized by a lowerReid vapor pressure and an increased octane number.

The alkylation reactor is generally maintained at a pressure (measuredat the feed inlet) in a range from 14 psia to 800 psia, optionally inthe range from 50 psia to 600 psia. The alkylation reactor is typicallymaintained at a temperature (measured at the alkylation reactor inlet)in a range from 150° C. to 350° C., optionally between 200° C. to 350°C. Typically, flow thorough the alkylation reactor is maintained at aweighted hourly space velocity (WHSV) in the range from 0.5 hr⁻¹ to 10hr⁻¹ on a hydrocarbon basis. Optionally, the WHSV is in the range from0.5 hr⁻¹ to 2.0 hr⁻¹. While higher overall throughput is desirable,ideally the chosen WHSV allows for conversion of at least 85% olefinicof hydrocarbons present in the mixed effluent at the selected operatingtemperature and pressure. The catalytic conversion occurring in thealkylation reactor produces an aromatic alkylation reactor effluent thattypically comprises at least 30 wt. % (preferably, at least 40 wt %) ofhydrocarbon molecules that are characterized by a boiling-point in therange of a liquid transportation fuel.

Speaking generally, the alkylation catalyst may comprise any catalystcharacterized as either Bronsted or Lewis acidic. A wide variety ofcatalysts have been found to promote aromatic alkylation including, butnot limited to, aluminum chloride, phosphoric acid, sulfuric acid,hydrofluoric acid, silica, alumina, sulfated zirconia, zeolites(including, for example, ZSM-5, ZSM-3, ZSM-4, ZSM-18, ZSM-20,zeolite-beta, H—Y, MCM-22, MCM-36 and MCM-49). In certain embodiments,the alkylation catalyst simultaneously promotes alkylation of aromaticsand oligomerization of olefins present in the mixed effluent.

Referring again to the embodiment depicted in FIG. 3, the alkylationeffluent 347 produced in the alkylation reactor 340 is conveyed to asecond separator 350 that separates the alkylation effluent 347 into twofractions: a light hydrocarbons fraction 351 comprising C1-C4hydrocarbons and H₂, and a heavy hydrocarbons fraction 352 comprisinghydrocarbons containing at least five carbon atoms (C5+) that may beutilized directly as a blend component of a liquid transportation fuelor additionally processed prior to blending into a liquid transportationfuel. Speaking generally, the alkylation effluent preferably comprisesan increased quantity (or increased wt %) of alkylated aromaticscontaining from seven to nine carbon atoms. Preferably, these alkylatedaromatics are monocyclic aromatic hydrocarbons.

Again, referring to the embodiment depicted in FIG. 3, lighthydrocarbons fraction 351 predominantly comprises hydrogen as well asC1-C4 hydrocarbons that remained unconverted in the alkylation reactor340. Light hydrocarbons fraction 351 leaves the second separator 350 andis optionally conveyed to a fourth separator 353 that utilizes aconventional separation technology (such as, but not limited to,pressure swing adsorption technology, membrane separation technology,etc.) to separate hydrogen from the light hydrocarbons to produce ahydrogen stream 354 and a light paraffins stream 355 comprisingparaffins continuing four or less carbon atoms. The light paraffinsstream 355 is optionally recycled to serve as a diluent during catalyticactivation of the first fraction 319. In such embodiments, the lightparaffins stream 355 is mixed with the first fraction 319 at a pointthat is upstream from the point where the first fraction 319 contactsthe activation catalyst 335 in the activation reactor 333.

The heavy hydrocarbons fraction 352 is conveyed to a third separator 360that is optionally a naphtha stabilizer. The third separator 360 isoperable to remove a unreacted C5/C6 fraction 372 (comprisingpredominantly alkanes and olefins containing four to six carbon atoms)from the heavy hydrocarbons fraction 352 in order to decrease the Reidvapor pressure and increase the octane rating of the resulting liquidhydrocarbon product 363 that predominantly comprises hydrocarbons(predominantly monocyclic aromatics) that are characterized by aboiling-point in the range of a liquid transportation fuel, such as, butnot limited to, gasoline, diesel and jet fuel. The unreacted C5/C6fraction 372 may optionally be used directly as a blend component 368 ofa liquid transportation fuel, or alternatively, is mixed withhydrocarbon feed stream 301 at a point that is upstream from theseparator 310. Optionally, at least a portion of the blending fraction325 and at least a portion of the liquid hydrocarbons fraction 338derived from the condenser/compressor 337 may be combined with theliquid hydrocarbon product 363.

Speaking more generally, in certain embodiments the liquid hydrocarbonproduct may be hydrotreated in a hydrotreating reactor containing ahydrotreating catalyst in order to reduce olefin and aromatic content inthe liquid hydrocarbon product, as well as to remove nitrogen-containingand sulfur-containing compounds. The hydrotreating reactor contains atleast one hydrotreating catalyst (such as, for example, NiMo, CoMo,etc.) or a precious metal catalyst (such as Pt/Al2O3, Pd/Al2O3, or Pd/C,etc) and is maintained at a pressure and temperature suitable forfacilitating hydrotreating catalytic reactions. Such processes areconventional in nature and therefore will not be described in greaterdetail here.

Certain embodiments comprise mixing a diluent with the first fractionprior to contacting the resulting mixture with an activation catalyst inthe activation reactor. The diluent may be added in a ratio ranging from10:1 to 1:10 molar ratio relative to the quantity of isomerizationeffluent fed to the activation reactor. The diluent may be added at anypoint that is upstream from, or within, the activation reactor, butprior to contacting the activation catalyst.

The diluent may comprise any substance that is less chemically-reactivewith the activation catalyst than the constituents that are present inthe first fraction at the conditions of temperature and pressure thatare maintained within the activation reactor. This prevents the diluentfrom reacting with the activation catalyst. Such properties are found ina large number of substances that are fully within the grasp of a personwho is knowledgeable in the field. In certain embodiments, the diluentmay comprise C1-C4 light paraffins, including recycling C1-C4 lightparaffins produced by the inventive processes and systems describedherein. In certain embodiments, the diluent may comprise any of methane,ethane, propane, butanes, benzene, toluene, xylenes, alkyl- ordialkyl-benzenes, naphthenes, C2-C5 olefins, and combinations thereof.

The presence of diluent during catalytic activation (i.e., activation)provides numerous advantages. First, it effectively decreases theconcentration of the isomerization effluent within the activationreactor. This results in a small increase in the total conversion ofalkanes to olefins or aromatics within the activation reactor. However,it also increases the selectivity toward the production of olefins,while slightly decreasing the selectivity toward the production ofaromatics. Adjusting the ratio of diluent to isomerization effluentchanges the ratio of olefins to aromatics in the resulting activationeffluent, thereby providing a valuable point of operational control fordownstream processes. Typically, the optimal molar production ratio ofolefins to aromatics ranges from about 0.5:1 to about 1.5:1, in order tomaximize the value captured in the olefin intermediates during thealkylation in the alkylation reactor. Mono-alkylated aromatics exhibitbeneficial (increased) octane rating and vapor pressure for applicationas blending components in certain transportation fuels such as gasoline.In contrast, di-alkyl and tri-alkyl aromatics comprising more than ninecarbon atoms are not well-suited for blending into gasoline, and exhibitnonoptimal cetane number for blending into diesel.

Addition of a diluent also advantageously favors the production ofvalue-added olefins relative to C1-C4 light paraffins and also mitigatesdimerization of C5 hydrocarbons to form durene(1,2,4,5-tetramethylbenzene), a problematic byproduct that oftenprecipitates from gasoline blends.

EXAMPLES

The following examples are representative of certain embodiments of theinventive processes and systems disclosed herein. However, the scope ofthe invention is not intended to be limited to the embodimentspecifically disclosed. Rather, the scope is intended to be as broad asis supported by the complete disclosure and the appending claims.

Example 1

This example demonstrates the preliminary rationale for isomerizing atleast a portion of the n-C5 found in a given hydrocarbon feed stream toi-C5 prior to contacting the pentane with an activation catalyst. Thegraphs below illustrate differences in activation reactivity for n-C5and i-C5. Feed streams were utilized that comprised either 100 wt. %i-C5 (i-C5), 100 wt. % n-C5 (n-C5), or a 70 wt. % iC5/30 wt. % n-C5mixture (i/n [7:3]) that represents an isomerized pentanes stream thathas been contacted with an isomerization catalyst to increase the ratioof i-C5 to n-C5. The catalyst utilized was ⅛″ extrudate consisting of 50wt. % alumina binder and 50 wt. % ZSM-5 zeolite, and experiments wereconducted at a WHSV of 1.3 hr⁻¹ at 1 atm. Results were averaged over thetotal time on stream of 16 hr.

FIG. 4 is a bar graph depicting the results of catalytically activatingeach feed stream at either 550° C. or 600° C. The graph depicts, aspercentages, the total catalytic conversion of each feed stream (firstcolumn), the yield to light olefins as product (second column), theyield to aromatics as product (third column) and the yield to C1-C4light paraffins (defined as non-olefin hydrocarbons containing from oneto four carbon atoms), fourth column. Yield was calculated on a wt. %basis, relative to the weight of the feed stream.

The results demonstrate that an effective conversion (82%) of a the i-C5fraction was achieved at 550° C., while a similar level of n-C5conversion (79%) required a temperature of 600° C. Activating the i-C5fraction at 550° C. (instead of 600° C.) also increased the yieldtowards the production of olefins while decreasing the aromatics yield.Lastly, these changes in yield caused no significant increase in theproduction of byproduct C1-C4 light paraffins. Meanwhile, theactivation/conversion of n-C5 at 550° C. only converted 48% of the feed,while the undesirable production of fuel gas (C1-C4 light paraffins) washigher compared to the same activation/conversion performed on a 100%i-C5 feed stream. The results for the 7:3 ratio feed of i-C5 to n-C5demonstrates the advantages of increasing the amount of i-C5 in the feedthat is fed to the activation reactor, as it shows many of the samebenefits demonstrated for the 100% i-C5 feed.

Example 2

This experiment demonstrates that isopentane (i-C5) is advantageouslyconverted by both an activation catalyst and a subsequentoligomerization catalyst to produce a high percentage of product that issuitable for use as a liquid transportation fuel. A 100 wt. % i-C5 feedstream was upgraded by first contacting it with a zeolite activationcatalyst, followed by contacting a zeolite oligomerization catalyst.Activation was conducted by contacting the feed stream with ⅛ in.diameter catalyst extrudate consisting of 50 wt. % alumina binder and 50wt. % ZSM-5 zeolite catalyst at a temperature of 579° C., and a WHSV of2.6 hr⁻¹ at 1 atm.

Oligomerization was conducted by contacting the activation effluent witha ZSM-5 zeolite catalyst in a reactor where the inlet temperature forthe activation effluent was maintained at 250° C., the pressure was 1atm, and the WHSV for the feed stream was 1.3 hr⁻¹. Results weretime-averaged over 16 hours. The table shows the product distributionfollowing conversion along with the selectivity to olefins and liquidproduct. The term “selectivity” indicates the percentage of thecatalytically converted feed stream that was converted to a particularproduct.

TABLE 2 Upgrading pentanes by activation alone or activation plusoligomerization. Activation + Activation Oligomerization TotalConversion (wt. %) 88 87 C1-C4 Light paraffins Yield 32 32 UpgradedProduct Yield (wt. %) 55 54 Total Coke Yield (wt. %) 0.1 0.1 LightOlefin Yield (wt. %) 42 16 Light Olefin Selectivity (wt. %) 48 19 Liq.Yield (wt. %) 13 38 Liq. Product Selectivity (wt. %) 15 44

The data in Table 2 show that the subjecting the effluent from the firstactivation reactor to a subsequent oligomerization step in a secondreactor increased the liquid product yield from 13 wt. % to 38 wt. %.This liquid product yield represents a liquid product suitable forblending into a liquid transportation fuel such as gasoline (up from 13wt. % prior to oligomerization), and that selectivity to liquid productfor the portion of the feed stream that was converted was 44 wt. %.Undesirable C1-C4 light paraffins production was limited to 32 wt. % ofthe original feed stream, which optionally may be recycled to be eitheractivated or to serve as a diluent in at least one of the activationreactors. Further, the final product only comprised 16 wt. % of lightolefins, (primarily ethylene), which may be recycled to the process, ordiverted to be utilized in any of a variety of conventional processes.

Example 3

This experiment demonstrates that isopentane (i-C5) is advantageouslyconverted by both an activation catalyst and a subsequent alkylationcatalyst to produce a high percentage of product that is suitable foruse as a liquid transportation fuel. A 100 wt. % i-C5 feed stream wasupgraded by first contacting it with a zeolite activation catalyst,followed by contacting the activation effluent with a zeolite alkylationcatalyst. Activation was conducted by contacting the feed stream with a⅛ in. diameter catalyst extrudate consisting of 50 wt. % alumina binderand 50 wt. % ZSM-5 zeolite catalyst in an activation reactor. Thetemperature of the activation reactor at the inlet for the feed streamwas 579° C., the pressure was 1 atm, and the WHSV for the feed streamwas 2.6 hr⁻¹. Alkylation was then conducted by contacting the effluentwith a ZSM-5 catalyst in a reactor where the temperature at the inletfor the feed stream was 230° C., the pressure was maintained at 15 psigand the WHSV of the feed stream was 1.3 hr⁻¹ at 1 atm. Results weretime-averaged over 16 hours. The table shows the product distributionfollowing conversion along with the selectivity to olefins and liquidproduct. The term “selectivity” indicates the percentage of thecatalytically converted feed stream that was converted to a particularproduct.

TABLE 3 Upgrading isopentane by activation only or activation followedby alkylation. Activation + Activation Alkylation Total Conversion (wt.%) 87 88.3 Light paraffins Yield (wt. %) 32.0 23.9 Light ParaffinsSelectivity 35.5 27.0 Upgraded Product Yield (wt. %) 55.0 64.4 TotalCoke Yield (wt. %) 0.1 0.2 Light Olefin Yield (wt. %) 42.0 14.4 LightOlefin Selectivity (wt. %) 48.0 16.3 Liquid Yield (wt. %) 13.0 50.0Liquid Product Selectivity 15.0 56.7

The data in Table 3 show that subjecting the activation effluent to asubsequent alkylation step increased the liquid product yield from 13.0wt. % to 50 wt. %. This liquid product is suitable for blending into aliquid transportation fuel such as gasoline, and possesses an increasedresearch octane number, a suitable distillation T50 and endpoint, andlow vapor pressure. Selectivity to liquid product for the portion of thefeed stream that was converted increased from 15.0 wt. % to 56.7 wt. %.Undesirable C1-C4 light paraffins production was limited to 23.9 wt. %of the feed stream, a portion of which could be used as diluent for theinventive processes and systems. Further, the final product onlycomprised 16.3 wt. % of light olefins. These olefins may be recycled tothe activation reactor, used as a diluent in the alkylation reactor, ordiverted to be utilized in any of a variety of conventional processes.

Note that the results shown in the above table may underestimate thetotal percentage of a mixed pentanes feed stream that would be availablefor blending into a liquid transportation fuel, as a typical hydrocarbonfeed stream (such as, but not limited to, natural gasoline) may alsoinclude an excess quantity of C5/C6+ that would not be eithercatalytically-cracked or introduced into the alkylation reactor, butwould partition into the second fraction and leave the system via anoutlet located upstream from the isomerization reactor. This excessquantity of C5/C6+ is suitable for direct blending into the liquidhydrocarbon product. In certain embodiments, a portion of the nC5/C6+fraction is diverted when necessary to achieve the desired 0.5:1 to1.5:1 olefin to aromatic ratio that maximizes production ofmono-alkylated aromatics in the alkylation reactor.

Example 4

Analysis of the product created in Example 3 by first activating, thenalkylating a 100 wt. % i-C5 feed stream revealed the composition shownin Table 4:

TABLE 4 Product composition and fuel characteristics for a 100 wt. %i-C5 feed stream that was activated at a temperature of 600° C. (WHSV =2.6 hr⁻¹) followed by alkylation at a temperature of 230° C. (WHSV 1.3hr⁻¹) Component Quantity (wt. %) Paraffin 3.0% iso-Paraffin 9.7%Aromatics 67.0% Mono-cyclic Aromatics 63.9% n-Olefins 1.1% iso-Olefins4.4% Research Octane Number (calculated) 98.0 Motor Octane Number(calculated) 89.6It is clear that the final product comprises a high wt. percentage ofmonocyclic aromatics that are an exceptional blend stock for a gasoline(see calculated Research Octane and Motor Octane Numbers, above). Only asmall quantity of paraffins remain, which can be recycled to process,used as diluent for the activation step, or combusted to produce atleast a portion of the heat required for the process.

Example 5

This example demonstrates the effect that a methane diluent has oncatalytic activation and conversion of two different hydrocarbon feedstreams: 1) a simulated “natural gasoline” comprising 50 wt. % i-C5 and50 wt. % n-C5 isomers, and 2) a simulated “first fraction” comprising 70wt. % i-C5 and 30 wt. % n-C5. Each of the two feed streams were fed at aWHSV of 1.3 hr⁻¹ to a reactor containing an activation catalystcomprising a ⅛″ extrudate consisting of 50 wt. % alumina binder and 50wt. % ZSM-5 zeolite. The temperature of the reactor (at the inlet forthe feed stream) was maintained at 600° C. and 20 psig (2.4 Bar) andresults were time-averaged for 16.5 hr. For certain reactions, methanediluent was co-fed along with each feed stream at a methane:feed streammolar ratio of 2:1.

The reaction produced an effluent comprising light olefins, aromaticsand light paraffins. Table 5 (below) shows the effect of the methanediluent on the total conversion of the 1:1 and 7:3 feed streams,respectively, as well as the selectivity of each conversion toward lightolefins, aromatics, and byproduct C1-C4 fuel gas.

TABLE 5 Catalytic activation of a 1:1 i-C5:n-C5 feed stream and a 7:3 1i-C5:n-C5 feed stream in both the absence and presence of methanediluent. Feed Stream 1:1 i-C5:n-C5 1:1 i-C5:n-C5 7:3 1 i-C5:n-C5 7:3 1i-C5:n-C5 +/− Diluent No Diluent CH₄ Diluent No Diluent CH₄ DiluentMaterial Balance 101%  103%  102%  101%  Conversion 92% 80% 93% 81% Fuelgas yield 37% 22% 34% 21% Product Yield 54% 58% 58% 60% Coke Yield  0% 0%  0%  0% Lt. Olefin Yield 34% 44% 35% 43% Lt. Olefin 37% 55% 38% 53%Selectivity Aromatic Yield 20% 14% 22% 17% Aromatic 21% 17% 24% 21%Selectivity Fuel Gas Yield 37% 22% 34% 21% Fuel Gas 41% 27% 37% 26%Selectivity

The data in Table 5 indicate that adding inert diluent caused slightloss of overall conversion, but significantly increased the yield andselectivity to light olefin production for both the 1:1 and 7:3 feedstreams. Adding inert diluent also greatly diminished selectivity toproduction of C1-C4 fuel gas. Meanwhile, only a small drop inselectivity to aromatics production was observed for the 1:1 ratio feedstream in the presence of diluent, which was offset by an equivalentincrease in aromatics production in the 7:3 ratio feed stream (in thepresence of diluent). All of these results are advantageous to theprocess, particularly in certain embodiments where the mixed effluent isimmediately utilized as feed stream for either an oligomerization oralkylation process. In certain embodiments that comprise anoligomerization process, diluent is added to the activation feed streamat a ratio that maximizes light olefin production, providing anadvantageous feed stream for the oligomerization catalyst. In certainembodiments that comprise an aromatic alkylation process, diluent can beadded to the activation feed stream at a ratio that produces a firsteffluent comprising olefins and aromatics at a ratio (typically between0.5:1 and 1.5:1 by mole) that provides an advantageous feed stream foran aromatic alkylation process.

In closing, it should be noted that the discussion of any reference isnot an admission that it is prior art to the present disclosure, inparticular, any reference that may have a publication date after thepriority date of this application. Although the systems and processesdescribed herein have been described in detail, it is understood thatvarious changes, substitutions, and alterations can be made withoutdeparting from the spirit and scope of the invention as defined by thefollowing claims.

Definitions

In the present disclosure, the term “conversion” is defined as any ofthe chemical reactions that occur during upgrading of hydrocarbons toliquid transportation fuels. Examples of such reactions include, but arenot limited to: oligomerization, aromatization, dehydrogenation,alkylation, hydrogenation and cracking.

We claim:
 1. A method for converting a feed stream comprising pentanesto produce a liquid transportation fuel, the method comprising: a)providing a hydrocarbon feed stream comprising at least 50 wt. %pentanes, including both n-pentane and isopentane; b) at least partiallyseparating the hydrocarbon feed stream in a separator to produce: afirst fraction that comprises at least 80% of the isopentane present inthe hydrocarbon feed stream (on a molar basis), and that furthercomprises at least 90% of hydrocarbons present in the hydrocarbon feedstream (on a molar basis) that are characterized by a vapor pressureequal to or greater than the vapor pressure of isopentane, and a secondfraction that comprises at least 80% of the n-pentane content of thehydrocarbon feed stream (on a molar basis) and at least 90% ofhydrocarbons content from the hydrocarbon feed stream that contains sixor more carbons (on a molar basis); c) contacting the first fractionwith an activation catalyst at a temperature and a pressure thatfacilitates catalytic activation of at least a portion of the firstfraction by the activation catalyst to produce an activation effluentcomprising olefins containing from two to five carbon atoms, monocyclicaromatics and unconverted alkanes containing from two to five carbonatoms; d) removing a first portion of the second fraction for use asblend component of a liquid transportation fuel; e) contacting a secondportion the second fraction with at least one isomerization catalyst inan isomerization reactor that is maintained at a temperature and apressure that facilitates the isomerization of at least a portion ofn-pentane in the second fraction to produce isopentane, therebyproducing an isomerization effluent characterized by an increased ratioof isopentane to n-pentane (on a wt. % basis) relative to the isopentaneto n-pentane ratio of the hydrocarbon feed stream; f) combining theisomerization effluent with the hydrocarbon feed stream; g) alkylatingat least a portion of the activation effluent of c) by contacting itwith an alkylation catalyst at conditions of temperature and pressurethat facilitates the alkylation of monocyclic aromatics to producealkylated monocyclic aromatics, thereby producing an alkylation effluentthat comprises an increased wt. % of mono-alkylated aromatichydrocarbons containing from seven to nine carbon atoms relative to themono-alkylated aromatic hydrocarbon content (in wt. %) of the activationeffluent.
 2. The method of claim 1, additionally comprising compressingand partially condensing the activation effluent to produce 1) a liquidhydrocarbons fraction comprising a portion of the monocyclic aromaticspresent in the activation effluent and unreacted alkanes containing atleast five carbon atoms, and 2) a compressed alkylation feed streamcomprising uncondensed monocyclic aromatics, hydrogen and C1-C5 alkanesand olefins, wherein the compressed alkylation feed stream is subjectedto the alkylating of g).
 3. The method of claim 1, further comprisingsplitting the alkylation effluent to produce a heavy hydrocarbonsfraction comprising at least 80 wt. % hydrocarbons that contain at leastfive carbon atoms, and a light hydrocarbons fraction comprising hydrogenand further comprising at least 80 wt. % hydrocarbons that contain fouror fewer carbon atoms.
 4. The method of claim 3, wherein the lighthydrocarbons fraction is separated to produce a hydrogen stream and alight paraffins stream comprising paraffins containing four or lesscarbon atoms.
 5. The method of claim 2, further comprising separatingthe liquid hydrocarbons fraction into an aromatics fraction and anunreacted C5/C6 hydrocarbons fraction, wherein the aromatics fractioncomprises monocyclic aromatics characterized as suitable for use as ablend component of gasoline, wherein the unreacted C5/C6 hydrocarbonsfraction comprises alkanes and olefins containing from five to sixcarbons.
 6. The method of claim 5, wherein the unreacted C5/C6hydrocarbons fraction is mixed with the hydrocarbon feed stream of a).7. The method of claim 1, wherein the hydrocarbon feed stream comprisesat least 70 wt. % pentanes.
 8. The method of claim 1, wherein theactivation catalyst comprises one or more zeolites characterized by aSi/Al ratio ranging from 12 to
 80. 9. The method of claim 1, wherein theactivation catalyst comprises ZSM-5 zeolite.
 10. The method of claim 1,wherein the activation catalyst facilitates carbon-hydrogen bondactivation that leads to at least one of olefination, dimerization,oligomerization and aromatization of the first fraction.
 11. The methodof claim 1, wherein the temperature in the activation reactor ismaintained at a temperature in the range from 500° C. to 625° C. and apressure in the range from 15 psig to 100 psig.
 12. The method of claim1, wherein the temperature in the activation reactor is maintained at atemperature in the range from 550° C. to 600° C. and a pressure in therange from 20 psig to 60 psig.
 13. The method of claim 1, furthercomprising adding a diluent to the first fraction prior to thecontacting of the first fraction with the activation catalyst, whereinthe diluent inhibits catalytic conversion of the first fraction by theactivation catalyst at the conditions of temperature and pressure thatare maintained in the activation reactor.
 14. The method of claim 1,further comprising adding a diluent to the first fraction prior to thecontacting of c), wherein the diluent is characterized as chemicallynon-reactive with the activation catalyst at the conditions oftemperature and pressure that are maintained in the activation reactor.15. The method of claim 13, wherein the diluent is added in an amountthat alters the specificity of the activation catalyst to increase theproduction of olefins, decrease the production of aromatics, orcombinations thereof, thereby increasing the ratio of olefins toaromatics in the activation effluent.
 16. The method of claim 13,wherein the diluent is added in an amount that is effective to producean activation effluent that is characterized by an olefins to aromaticsratio in the range of 0.5 to 2.0.
 17. The method of claim 13, whereinthe diluent is selected from at least one of methane, ethane, propaneand butane isomers.
 18. The method of claim 13, wherein at least aportion of the light paraffins stream is mixed with the first fractionand serves as the diluent.
 19. The method of claim 1, wherein at least aportion of the second fraction is passed through multiple isomerizationreactors that are arranged in series configuration, wherein each of themultiple isomerization reactors contains at least one isomerizationcatalyst and is maintained at a temperature and a pressure thatfacilitates the isomerization of n-pentane to isopentane by the at leastone isomerization catalyst.